Process for manufacturing highly concentrated hydroxylamine

ABSTRACT

An improved ion exchange process for manufacturing and concentrating hydroxylamine from an aqueous solution comprising hydroxylammonium ion and counter anions is disclosed. The process involves the use of an aqueous hydroxylamine wash step which assists in controlling the processing temperature and leads to a more concentrated, purified aqueous hydroxylamine solution.

BACKGROUND OF THE INVENTION

1. Field of the Invention

This invention relates to a process for producing hydroxylamine, andmore particularly, to such process utilizing an ion exchange resin toproduce purified, concentrated solutions of hydroxylamine free base.

2. Description of the Related Art

Hydroxylamine is prepared commercially by the Raschig process orvariations thereof in which ammonium or sodium nitrite is reacted inaqueous solution with ammonium or sodium bi-sulfite and sulfur dioxideand the resulting disulfonate salts are subsequently hydrolyzed to asolution containing essentially hydroxylammonium sulfate, sulfuric acid,ammonium sulfate and/or sodium sulfate plus minor amounts of thecorresponding nitrates. This solution can be used, after neutralizationwith ammonia, as a source of hydroxylamine or pure hydroxylammoniumsalts from the mixture.

One method of obtaining pure hydroxylammonium salts consists of using ahydroxylammonium containing mixture to synthesize an oxime from aketone, separating the oxime from the spent solution and hydrolyzingthis oxime with a strong mineral acid to recover hydroxylammonium saltand the ketone which can be recycled. This method uses long periods ofheating for the hydrolysis and requires expensive equipment for theseparation of the oxime from the spent solution and of thehydroxylammonium salt from the ketone. Moreover, salts of hydroxylaminewith weak or oxidizing acids cannot be prepared by this method becausethese acids either do not effect hydrolysis of oximes or decompose thehydroxylammonium salt formed during hydrolysis. Salts of such acids canbe prepared by neutralizing cold solutions of hydroxylamine with thecorresponding acid.

Heinz Holsapfel in Z. Anorg. und Allgem. Chemie, Vol. 288, page 28(1956) describes the preparation of hydroxylamine from hydroxylammoniumsalts by employing an anion exchange resin in the OH form. Thispublication does not relate to the separation of hydroxylamine fromsolutions containing other cations. As disclosed in U.S. Pat. No.3,508,864 issued Apr. 28, 1970 to Wallace T. Thompson et al.,hydroxylammonium perchlorate can be produced either by liberatinghydroxylamine from a hydroxylammonium salt by passage through an anionexchange resin and neutralizing with perchloric acid, or by absorbinghydroxylammonium ion an a cation exchange resin and then passingperchloric acid through the resin. This method does not separatehydroxylamine from other cations. A cation exchange method is describedby Earl. J. Wheelwright in Industrial Engineering Chemistry ProcessDesign Development, Vol. 16, No. 2 (1977), page 220 for the preparationof hydroxylammonium nitrate. This method does not separate hydroxylaminefrom other cations and moreover the resulting solution ofhydroxylammonium nitrate contains significant amounts of nitric acid.

Other methods for manufacturing, recovering, and concentratinghydroxylamine solutions and hydroxylammonium nitrate are known in theart. U.S. Pat. No. 5,213,784 discloses a process for producingconcentrated purified hydroxylammonium nitrate. The process involvesadmixing nitric acid having a concentration less than about 70% tosolutions containing excess hydroxylamine. The process avoidsspontaneous decomposition of the product.

U.S. Pat. No. 4,725,360 discloses a process for recovering hydroxylaminefrom waste water in the form of hydroxylammonium sulfate. The processconsists of passing the waste water over a strongly acidic ion exchangerand then eluting hydroxylamine associated with the ion exchanger withsulfuric acid.

U.S. Pat. No. 4,202,765 discloses a method for purifying hydroxylamineusing anion exchanger. The process consists of passing a hydroxylaminecontaining solution across a bed of cationic exchange resin anddesorbing absorbed hydroxylamine using a monovalent amine or hydroxidebase.

The prior art describes methods for preparing hydroxylamine andhydroxylammonium salts, but there is still a need for a simple,inexpensive, and continuous process for separating hydroxylamine fromsolutions containing ionic contaminants. There is also a great need fora process that is capable of producing a hydroxylamine solutionscontaining essentially no anion impurities such as sulfates.

SUMMARY OF THE INVENTION

This invention relates to a process for producing hydroxylamine, andmore particularly, to such process utilizing ion exchange. Additionally,the process employs various processing steps whereby a concentratedsolution of hydroxylamine is obtained in a purified state.

DESCRIPTION OF THE DRAWING

The present invention will be more readily understood by reference tothe following drawing taken in conjunction with the detaileddescription, wherein:

FIG. 1 is a graphical representation of effluent concentrations ofhydroxylamine (HA) and NH₃ using the present invention with a 1-inch ionexchange column;

FIG. 2 is a graphical representation of hydroxylamine (HA) productconcentration obtained with the present invention with time whichreflects the effect of a recycled HA aqueous solution step;

FIG. 3 is a graphical representation of the internal column temperaturewith time during the production of HA, using the present invention, withand without using a recycled HA solution;

FIG. 4 is a graphical representation of effluent concentration profilesof HA production using the present invention;

FIG. 5 is a graphical representation illustrating the effect of CDTAstabilizer on the self-heat rate of a 57% by weight HA solution.

DETAILED DESCRIPTION OF THE INVENTION

This present invention rates to an ion-exchange process formanufacturing hydroxylamine from an aqueous solution includinghydroxylammonium ion and ionic contaminants such as H⁺, NH₄ ⁺, metal andsulfate ions.

The ion-exchange process of this invention is useful for producing ahigh purity aqueous hydroxylamine solution for subsequent manufacture ofhydroxylamine salts such as hydroxylamine nitrate as well ashydroxylamine sulfate, and other hydroxylamine salts for variousintermediate and end uses. An additional advantage of the process ofthis invention is its ability to produce a purified aqueoushydroxylamine product that is essentially free of cationic contaminantsas well as anionic contaminants such as sulfate ion. The term“essentially free” as used herein, refers to a purified aqueoushydroxylamine product that has less than 400 ppm anionic contaminantssuch as sulfates and preferably 200 ppm or less of anionic contaminants.The term “essentially free” when used to describe ionic contaminantsmeans less than 200 ppb ionic contaminants including all metal cations.

The ion exchange process of this invention is useful for purifying andconcentrating hydroxylamine from hydroxylammonium ions produced in avariety of processes. Such processes are disclosed and described, forexample, in U.S. Pat. Nos. 5,213,784 and 4,491,567, which describeprocesses for producing hydroxylammonium ion salts such ashydroxylammonium nitrate (HAN) and hydroxylammonium sulfate (HAS), andwhich are incorporated herein by reference in their entirety. A commonprocess for preparing a hydroxylammonium ion containing solution is theRaschig process. The Raschig process produces an aqueous solutioncomprising hydroxylammonium ions, hydrogen ions, ammonium ions, andsulfates.

The process of this invention is useful for producing a purified andconcentrated aqueous hydroxylamine solution from any aqueoushydroxylammonium ion containing solution, including, but not limited to,Raschig solutions, HAN solutions and HAS solutions. However, the processof this invention will generally be described in the context ofpurifying and concentrating hydroxylamine from the aqueoushydroxylammonium ion containing solution product of the Raschig process,and HAS solutions.

The process of this invention uses an ion-exchange resin in a container,e.g. in column form, loaded with a cation exchange resin to bind to thehydroxylammonium ion in order to accomplish hydroxylamine purificationand concentration. A suitable cation exchange resin is selected. Alltypes of cation exchange resins are suitable, e.g., sulfonic, phosphoricor carboxylic types. Sulfonic resins are preferred because they have ahigh selectivity for hydroxylammonium ion over other cations and anions.

In this inventive ion-exchange process, desorption of thehydroxylammonium ion from the resin which is contained in the exchangeresin must be accomplished. This is typically accomplished by employinga suitable base solution, e.g. sodium hydroxide, potassium hydroxide,ammonium hydroxide, etc., that has a higher affinity towards thecationic exchange resin than hydroxylamine.

For displacing hydroxylammonium ion from the resin, any water solublemonovalent amine or hydroxide base can be used. Polyvalent bases areless suitable because their cations have great affinity for the exchangesites of the cationic resin and therefore cannot be readily displaced byhydroxylammonium ion in a subsequent ion-exchange loading cycle.Suitable monovalent bases include, for example, sodium hydroxide,potassium hydroxide, lithium hydroxide and the like; methylamine,ethylamine, dimethylamine, diethylamine, and the like. The preferredbase is ammonium hydroxide which has the advantages of being relativelyinexpensive, non-toxic, of low molecular weight, and volatile. Moreover,the by-product obtained when ammonia is used as a base for the recoveryof hydroxylamine from a mixture of sulfate salts is ammonium sulfate,which is valuable as a fertilizer. A most preferred desorbent base is anaqueous ammonia solution having an ammonia concentration of at least 30%and preferably a concentration of at least 40%.

The volatility of ammonia allows for the use of a slight excess ofammonia in the desorption step in order to displace hydroxylaminecompletely from the resin because small amounts of ammonia that willcontaminate the product hydroxylamine (less than about 3% during thisoperation can easily be removed by evaporation. Furthermore, any productcontaminated with an unacceptable level of ammonium ion can be recycledto feed the stream. It has been found, however, that it is possible toobtain essentially ammonia-free and ionic contaminant free hydroxylaminesolutions directly from the resin using a deficient amount of ammonia todisplace hydroxylamine.

The aqueous base solution which is used to desorb hydroxylamine from theion exchange resin during the desorption step may optionally containhydroxylamine. When the aqueous base solution includes hydroxylamine,the ion exchange column effluent which is thereafter finally obtained bysubsequently rinsing the resin bed with water, contains hydroxylamine ata concentration which is substantially higher than the concentrationobtain if no hydroxylamine were present in the aqueous base solution.This effect allows the recovery of relatively concentrated hydroxylamineliquors in a cyclic and/or continuous ion exchange process by simplyrecycling part of the hydroxylamine product solution to the unloadingstep of the next cycle after mixing with a base.

The ion exchange process of this invention is conducted in essentiallyfive steps. The first two steps are defined as the “loading steps” andconsist of: (1) an ion exchange step in which a hydroxylammonium ioncontaining solution is passed across the cation exchange resin in orderto allow the hydroxylammonium ion to bind to the cation exchange resin;and (2) a wash step whereby deionized water or some other wash solutionis directed across the cation exchange resin in order to maximize theassociation of hydroxylammonium ion with the cation exchange resin. Thewater wash step also removes any non-exchanged ionic contaminants suchas hydrogen ions, ammonium ions, and anionic contaminants, such asanionic sulfates, and metal ions such as calcium ions, sodium ions,ferrous and ferric ions, etc., from the solution filling the void spacesbetween the cation exchange resin particles, i.e. becomes “associated”with the resin. In addition, the wash water can also be used to adjustthe temperature of the resin column before the subsequent desorptionstep.

The next step is a temperature maintenance step and a concentrationenhancing step which comprises a wash of the resin column with anaqueous solution of hydroxylamine (HA) to displace at least a portion ofwater from the pores of the resin and in the void space between resinparticles. It has been found that in order to get pure HA aqueoussolutions in concentrations greater than 12–13%, the use of the recycleHA and introduction of a highly concentrated desorbent are essential.The use of highly concentrated desorbent will add to the desorptiontemperature due to the heat of dilution of the desorbent. The use ofrecycle HA will keep the processing temperatures, i.e. the temperatureof the resin column, at a maximum of 50° C. to 80° C. throughout theentire process. If this temperature is maintained, the thermaldecomposition of the resultant HA is minimized and the resultant pure HAproduct is obtained in an excess of 50% weight percent aqueous solution.

Additionally, it is to be noted that the temperature control, e.g.50°–80° C., prevents decomposition of the resultant HA, thereby leadingto a more pure product. Finally, the utilization of the aqueous HArecycle step maintains the process temperature when the subsequentdesorption of the hydroxylammonium ion from the resin is conducted, asdiscussed hereafter. Such desorption involves an exothermic heat ofexchange which is thus modulated by the HA aqueous recycle.

It is to be noted that the aqueous solution of HA is typically obtainedby recycling a portion of the hydroxylamine product after the unloadingsteps hereafter described.

The final step of the process are known as the “unloading steps”. Theunloading steps are: (4) a desorption step; and (5) a rinse step. Aspreviously discussed, the desorption step uses concentrated aqueous basesolution comprising a soluble amine or hydroxide base that has a higheraffinity towards the cationic exchange resin than the hydroxylammoniumion. The soluble amine or hydroxide base replaces hydroxylammonium ionan the cationic resin thereby desorbing hydroxylamine and making itavailable for collecting in the ion exchange column effluent stream. Arinse step follows the desorption step and uses deionized water or someother aqueous rinse solution to elute the desorbed hydroxylamine fromthe ion exchange column where it is collected in a purified andconcentrated product. The rinse step also cleans the resin bedsufficiently so that no liquid phase reactions occur between thedesorbent or eluted products and the subsequent feed.

It is to be stressed that unlike previous ion exchange processes, theresultant aqueous HA product contains metallic ion impurities in therange of about 10–20 ppb or below, as compared to other processes whichyield a product containing about 5 ppm. And, as previously indicated,unlike other ion exchange processes, a concentration of such product isobtained in the 50% by weight concentration and higher range as comparedto a yield of 12–13% by weight concentration.

Each step is described in greater detail below.

The Ion Exchange Step

A feed solution comprising hydroxylammonium ion and ionic impuritiessuch as sulfate or nitrate impurities is fed into and through a columnof cation exchange resin during the ion exchange step. The feed solutionmay include other ionic contaminants such as hydrogen ion, ammonium ion,metal ions, sulfate ion, nitrate ion, chloride ion, and phosphate ion.The type and amount of feed contaminants present will depend largely onthe process used to manufacture the hydroxylammonium ion containingfeed. However, typically, an aqueous feed solution comprising 20 to 40weight percent of a hydroxylammonium salt, e.g. a 30 to 35 weightpercent hydroxylammonium sulfate (HAS) solution is fed into and throughthe resin column to form a hydroxylamine loaded resin.

The hydroxylammonium ion fed into the resin packed column is retained bythe cation exchange resin until the resin is at least partiallysaturated with hydroxylammonium ion at which time the hydroxylammoniumion begins to appear in the column effluent and breakthrough is reached.This first method of loading the resin with hydroxylammonium ion untilthe breakthrough point is reached represents the maximum amount ofhydroxylammonium ions that may be quantitatively removed by the cationicexchange resin from the feed solution while minimizing loss ofhydroxylammonium ion in the exchange step column effluent. Continuinghydroxylammonium ion feed into the resin packed column after thehydroxylamine breakthrough occurs represents a second method of resinloading whereby hydroxylammonium ion appear in the column effluentstream in progressively larger concentrations until the mole-fraction ofhydroxylammonium ion in the effluent approaches the mole-fraction ofhydroxylammonium ion in the feed solution. The resin of at this point isessentially at equilibrium with the feed solution and thehydroxylammonium ion loading an the cation exchange resin represents theequilibrium capacity of the resin for hydroxylammonium ion in thepresence of excess feed solution.

The second method for loading the cation exchange resin withhydroxylammonium ion described above maximizes the ability of thecationic exchange resin to load hydroxylammonium ion but it produces anunpurified effluent that contains large amounts of desiredhydroxylammonium ion. This stream can be recycled to the feed stream ifnecessary.

It is preferred that the concentration of hydroxylammonium ion in thecolumn effluent during the ion exchange step be minimized whilemaximizing the loading of hydroxylammonium ion on the cationic exchangeresin. This is accomplished by carefully monitoring various parametersof the effluent stream emanating from the ion exchange column during theion exchange step as described in U.S. Pat. No. 5,762,897, incorporatedherein by reference in its entirety.

As previously discussed, the purity and concentration of the resultantHA product is dependent on the temperature of the resin columnmaintained throughout the process. Hydroxylamine product containingabout 10–20 ppb or less of metallic ion impurity can only be obtained ifthe ion-exchange resin column is maintained at a temperature of 40 to100° C., preferably about 50° to about 70° C.

The feed linear velocity to and through the ion exchange resin column istypically 2.0 to 5.0 cm/min. If the feed linear is not maintained atthis rate the maximum loading of the hydroxylammonium ion an the resinmaterial may not be achieved.

The cation exchange resin that is loaded in an ion exchange columnuseful in the process of this invention will typically be in particulateform of a granule or a spherical bead. As a result, there are spacesbetween particles that are occupied or associated by any fluid fed intothe ion exchange column during the various process steps. During the ionexchange step, the hydroxylammonium ion, anionic contaminants andcationic contaminants containing feed occupies the void spaces betweenthe cation exchange resin particles and must be flushed from the columnprior to desorbing hydroxylammonium ion from the resin. This is done inorder to prevent anionic contaminants and other feed contaminants fromcontaminating the essentially pure aqueous hydroxylamine product.

Accordingly, the resultant hydroxylammonium ion loaded resin is treatedwith wash water in a washing step to remove any excess ofhydroxylammonium ions, e.g. HAS solution and the ionic impurities(anionic and cationic), from the resin to form a purifiedhydroxylammonium ion loaded resin. Such water washing is typicallyconducted by the process described in U.S. Pat. No. 5,762,897.

The Temperature Maintenance and Concentration Enhancing Wash

As previously indicated, the temperature of the ion-exchange column mustbe controlled because of the exothermic heat of exchange which occursduring the ultimate—desorption of hydroxylamine by a base solution. Thisis accomplished by washing the resultant purified hydroxylammonium ionloaded resin with the aqueous hydroxylamine (HA) solution. The aqueousHA solution is a low concentration solution, typically 25–34 percent byweight of HA.

The HA solution is fed into the ion exchange resin, typically at alinear velocity of 2.0 to 5.0 cm/min, to replace the water contained orassociated in the void spaces between the cation exchange resinparticles to form a HA loaded resin. The HA wash solution is typicallyrecycled from the resultant HA product after the desorption and rinsestep.

Again, it is to be stressed that this process step produces a lowertemperature during the subsequent desorption step, which in-turn reducesthe decomposition of the HA during its production and ultimately leadsto a very pure HA product.

The Desorption Step

The purpose of the desorption step is to feed a desorbent solution intothe ion exchange column that includes ions that will preferentiallyreplace hydroxylammonium ion at the cationic exchange resin bindingsites thereby eluting hydroxylamine from the resin and from the column.As discussed above, useful desorbents are preferably monovalent aminebases, hydroxide bases, or any combination thereof.

A preferred desorbent is aqueous ammonium hydroxide having aconcentration of from about 40 to about 55% by weight. It is mostpreferred that the aqueous ammonium hydroxide solution have aconcentration of at least 45%. The aqueous ammonium hydroxide solutionis added to the aqueous HA solution washed ion exchange column at alinear velocity of from 1.5 to 3.5 cm/min. It is important to maintainthis low linear velocity of the desorbent solution in order to ensurethat hydroxylamine is efficiently and completely desorbed from thecation exchange resin. It is also preferred that a predetermined amountof desorbent be added to the HA solution washed ion exchange column. Theamount used will depend upon the ion exchange resin used, the desorbent,and upon the desorbent concentration. For example, where a 48% by weightammonia solution of aqueous ammonium hydroxide is used as a desorbent,an amount of desorbent equal to 4.5 to 4.7 milliequivalents of ammoniumhydroxide per milliliter of resin should be added during the desorptionstep. The temperature of the desorbent stream can also be adjusted forthermal control of the heat generated during desorption.

At the end of the desorption step, the ion exchange column containsenough desorbent material to desorb bound hydroxylammonium ion from thecation exchange resin. However, the volume of desorbent added to the ionexchange column is generally very small and hydroxylamine does nottypically appear in the ion exchange column effluent stream whiledesorbent is added to the ion exchange column. Therefore, a rinse stepis necessary to propel the desorbent through the column so as tocompletely desorb hydroxylamine from the cation exchange resin and toprepare the cation exchange resin for a subsequent ion exchange step.

The Rinse Step

The rinse step performs two functions. It moves hydroxylamine out of thecolumn and into the effluent stream from the ion exchange column and itprepares the ion exchange column for a subsequent sequence of ionexchange steps. The most important aspect of controlling the rinse stepis to determine when to begin collecting a hydroxylamine containing ionexchange column effluent stream product and when to halt collecting thehydroxylamine containing ion exchange effluent stream product in orderto maximize the volume and concentration of the hydroxylamine productwhile minimizing any contaminants in the collected product. It has beendiscovered that by closely controlling rinse parameters, such as rinsewater temperature and linear velocity, and by monitoring the ionexchange column effluent stream parameters such as density, conductivityand pH, the purity can be maximized and the concentration ofhydroxylamine in the effluent stream emanating from the ion exchangecolumn during the rinse step can also be maximized.

It is preferred that the rinse step be performed using deionized waterhaving a temperature from about 15° to about 30° C. The warm waterincreases the kinetics and solubility of unused desorbent and improvesdesorption kinetics as well. The linear velocity of the rinse water fedto the ion exchange column during the rinse step should be maintained atfrom about 2.0 to about 5.0 cm/min.

It is preferred that the effluent stream from the exchange column duringthe rinse step be collected as soon as hydroxylamine is detected in theeffluent stream. This occurs by measuring the density, conductivity ofthe effluent and/or the ion exchange column bed height and/or theeffluent stream pH, as described in U.S. Pat. No. 5,762,897.

The cation exchange resin may be employed in the process in the form ofa dense compact fixed bed which is alternatively contacted with the feedmixture and desorbent materials. In the simplest embodiment of theinvention, the ion exchange resin is used in the form of a single staticbed in a semi-continuous process. In another embodiment, two or more andpreferably four static ion exchange resin beds are used in associationwith appropriate valving so that the feed solution may be passed throughone or more ion exchange resin containing beds while the desorbentsolution is passed through one or more of the remaining static beds. Theflow of feed solution and desorbent solution may be either up or downthrough the resin bed. Furthermore, any conventional apparatus employedin static bed fluid-solid contacting may be used to accomplish theprocess of this invention.

Counter-current moving-bed or simulated moving-bed counter-current flowsystems are preferably used in the process of this invention becausethey have a much greater separation efficiency than fixed adsorbent bedsystems. In the moving-bed or simulated moving-bed process, theadsorption and desorption operations are continuously taking place whichallows both continuous production of an extract and a raffinate streamand the continual use of feed and desorbent streams. The raffinatestream comprises feed impurities, desorbent, and so forth and iscomparable to the effluent from the ion exchange step described aboveand to the water wash step effluent stream.

One preferred embodiment of this process utilizes what is known in theart as the simulated moving-bed counter-current flow system, asgenerally described in U.S. Pat. No. 5,762,897, with modification as tothe number of zones.

Where the resultant HA product contains NH₃ contamination, e.g. 2–3weight percent of NH₃, the HA product is treated by stripping the NH₃from the solution. The stripping can be achieved by purging withnitrogen or by a vacuum stripping, e.g. at 35°–45° C. at a vacuum of60–15 mm Hg.

For applications in the semi-conductor industry that requires metalimpurities at <5 ppb, an after-treatment process is employed.Conventional methods for the removal of metal ions utilizecation-exchange resins. However, to achieve a high purity product,several types of cation-exchange resins may be required and multi-stepprocesses are used. It has been discovered that metal chelating agents,such as CDTA (cyclohexanediaminotetraacetic acid) and others, areeffective for the complexation of metal ions in the presence of highlyconcentrated HA solutions. Furthermore, the chelating agents arenegatively charged.

Therefore, these complexes and metal ions can be removed from the HAsolution with an anion-exchange resin. The general method can be appliedto the purification of coolants, such as propylene glycol, that areutilized for cooling HA process streams for the safe operation of the HAprocess.

In addition, a NH₄-form of common strong acid cation exchange resin,e.g. Rohm & Haas Amberjet 1500, is used for the removal of alkalinemetals, such as Na and K, in the presence of highly concentrated HAwithout the difficulty of using H+-form of cation exchange resin thatgenerates a lot of heat causing the fast decomposition of HA.

EXAMPLES Examples 1 and 2

Using a 1-inch ID×37.7″ long (PFA) column with 485 cm³ of Rohm and HaasAmberjet brand 1500 resin, Examples 1 and 2 illustrate the invention.The resin was pretreated with 40% sulfuric acid to remove metals on theresin and exchanged to NH₄ ⁺-form using a 14% aqueous NH₃ solution. 35%by weight aqueous HAS feed was pumped into the column at a flow rate of16 cm³/min for 14 minutes. This was followed by a water wash at a flowrate of 17 cm³/min for 54 minutes. At the end of the water wash, arecycled HA aqueous solution (34% by weight) was fed into the column at16 cm³/min for 8 minutes and followed by the introduction of an aqueousNH₃ (45% by weight) solution at a flow rate of 17 cm³/min for 5 minutes.A water rinse was followed immediately at a flow rate of 16 cm³/min for53 minutes. Effluent stream samples were collected every 2 minutes.Concentrations of HA and NH₃ in the product recovery region were plottedagainst bed volume as illustrated in FIG. 1. As a comparison, a test runwas conducted under the general procedure described above except that noaddition of the recycled HA solution was employed. Effluent productconcentrations with and without recycled HA were compared as shown inFIG. 2. A decrease in the internal column temperature with the recycledHA solution is illustrated in FIG. 3.

As shown in FIG. 1, pure products of 48–50% by weight HA and greater areobtained without the need of a concentration step(s).

Examples 3–6

Following the general procedure described in Examples 1 and 2, a 4-inchPFA-lined column was used. The HAS feed flow rate was at 240 cm³/min,wash and rinse water at 240 cm³/min, recycled HA at 220–240 cm³/min andaqueous NH₃ at 160–220 cm³/min. The HAS was fed for 16 minutes whilewash and rinse water were 52 and 50 minutes, respectively. Scales wereused to monitor the weights used for the recycled HA and the aqueousNH₃. Parameters used for these examples are summarized in Table 1,below.

TABLE 1 HA Product Yield Efficiency with a Single 4-inch PFA-linedColumn Example 3 4 5 6 Resin volume (cm³) 8617 8770 8641 8641 HA loaded(g) 568.7 578.8 570.3 570.3 Recycle HA (g) at 34% 2240 2220 2210 2200Aqueous NH₃ (g) at 50% 1330 1361 1330 1283 Product (49%), cm³ 1006.6980.85 1017.2 910.2 Gram 1115.0 1066.5 1114.3 993.3 Net HA (g) 546.4522.6 546.0 486.7 Recoverable (40% HA, 433.0 538.1 375.0 767.9 2–3%NH₃), cm³ Gram 459.8 501.3 416.4 777.3 Net HA (g) 183.9 200.0 166.5310.1 Recycle HA, (23–34%) cm³ 1438.5 1653.6 1200.1 1079.1 Gram 1526.61718.5 1284.5 1141.2 Net HA (g) 458.0 378.1 436.7 365.2 TotalRecoverable Ha (g) 1188.3 1100.7 1149.2 1162.0 Net HA produced (g) 242.8145.9 231.3 103.9 Efficiency (%) 42.7 25.2 40.6 18.2

Effluent stream samples were collected every two minutes for chemicalidentification and analysis while products were collected every minuteto determine product concentration and yield. Products collected for thefour examples are tabulated in Table 2 below while the concentrations ofHA and NH₃ in the product samples are illustrated in FIG. 4. From theproduct concentration profile, samples of pure HA averaging 48–50% byweight were combined as the product cut. Product samples that containpure HA with averaging concentrations at 22–34% were collected as therecycle HA cut. And lastly samples that average HA at 40% and NH₃ at2–3% by weight were combined as the potentially HA recoverable product.

TABLE 2 Product Collected Per Minute (cm³) Product Collected @ Exampletime (min) 3 4 5 6 95 242.11 207.52 96 243.71 246.28 199.25 97 217.76251.64 230.68 206.57 98 207.85 244.97 232.28 227.4 99 176.98 245.56224.85 238.34 100 156.3 231.02 211.5 230.26 101 193.79 203.6 205.37225.26 102 192.55 230.57 193.93 228.1 103 199.95 230.79 192.96 226.54104 213.93 245.4 192.96 263.2 105 187.48 258.68 203.36 243.5 106 212.71245.98 203.89 261.24 107 207.34 224.63 209.82 253.67 108 225.66 313.42199.15 109 197.18 110 207.96 111 186.15 112 188.8

One objective of this invention is a process to manufacture aconcentrated hydroxylamine free base with high purity. Table 5, below,summarizes results in parts per billion (ppb) with (ICP-AES) analysis ofthe starting feed and the HA product produced in Examples 3–6. The tableillustrates the capability of this invention to manufacture concentratedHA with ppb level metals from a HAS feed that contains metals at about 1ppm level.

TABLE 5 Production of a High Purity HA with this Invention Metal contentHAS (ppb) feed Example 3 Example 4 Example 5 Example 6 Ca 710 <10 <10 <10  <10 Fe 850 <10 52 73  70 Na 1200  <10 60 11 <10

Examples 7–13

The recoverable product cut that contains a HA concentration of 38–42%by weight and NH₃-concentration of 2–3% by weight was transferred to a2-liter rotary evaporator. For these examples, about 1200–1400 grams ofthe solution were first added with 70 ppm of CDTA stabilizer andtransferred to the evaporator. A mechanical pump was used to evacuatethe system while a condenser in the evaporator was cooled with a watercoolant at a temperature of about 4° C. At the end at the strippingprocess, the concentration of HA was also increased somewhat. Results ofthe stripping tests are tabulated in Table 3, below. Combining productsfrom both the product cut and the HA recovered from the strippingprocess, the net product efficiencies increase substantially as shown inTable 4, below.

TABLE 3 Vacuum stripping of NH₃ from the product cut that has NH₃contamination Initial HA Initial NH₃ Initial Final HA Final wt Total wtExample # (wt %) (wt %) wt (g) Net HA (g) (g) (g) (g)  7 41.89 3.13 1173491.5 53.97 882 881.5  8 41.89 3.13 1235 517.22 53.45 848 1729.7  939.48 1.97 1350 532.98 55.83 733 4451.5 10 39.48 1.97 1239 489.32 52.67760 5211.1 11 39.24 2.696 1350 529.7 47.91 1060  9732 12 39.24 2.6961163 456.17 50.07 835 10567 13 39.24 2.696 1229 482.34 49.33 913 11479Final Net Condensate Condensate Condensate HA Recovery Tem. Vacuum TimeExample # HA (g) HA (%) NH₃ (%) (wt) (%) (° C.) (torr) (min)  7 475.7513.63 1.193 253.8 103.83   37–38.5   66–12.7  90  8 453.36 13.63 1.193213.5 93.28 37–38   62–12.9  90  9 409.46 12.61 1.106 412.2 86.58 37–38  62–12.8 140 10 400.08 12.61 1.106 382.8 91.63 37–38   52–12.9 120 11507.8 9.65 1.2 346.7 102.18 38–39 70–13 120 12 417.88 11.5 1.2 279.798.66 38.5 79–13 120 13 450.19 12.2 1.2 264.5 100.02   35–38.5 70–13 120

TABLE 4 Enhanced HA Product Yield Efficiency with a Single 4-inchPFA-lined Column Example 3 4 5 6 Resin volume (cm³) 8617 8770 8641 8641HA loaded (g) 568.7 578.8 570.3 570.3 Recycle HA (g) at 34% 2240 22202210 2200 Aqueous NH₃ (g) at 50% 1330 1361 1330 1283 Product (49%), cm³1006.6 980.85 1017.2 910.2 Gram 1115.0 1066.5 1114.3 993.3 Net HA (g)546.4 522.6 546.0 486.7 Recoverable (40% HA, 433.0 538.1 375.0 767.92–3% NH₃), cm³ Gram 459.8 501.3 416.4 777.3 Net HA (g) 183.9 200.0 166.5310.1 Recycle HA, (23–34%) cm³ 1438.5 1653.6 1200.1 1079.1 Gram 1526.61718.5 1284.5 1141.2 Net HA (g) 458.0 378.1 436.7 365.2 TotalRecoverable Ha (g) 1188.3 1100.7 1149.2 1162.0 Net HA produced (g) 426.7345.9 397.8 414.0 Efficiency (%) 75.0 59.8 69.8 72.6

Examples 14–15

Hydroxylamine solutions manufactured according to the above-mentionedembodiments were treated with a cation-exchange resin in NH₄ ⁺-form(Rohm & Haas Amberjet 1500). The HA solution was stabilized with CDTA ata concentration level of 100 ppm. This solution was pumped at a flowrate of 15 cm³/min into a 1-inch column with 300 cm³ of anion-exchangeresin (Rohm and Haas Amberjet 4200). The resin was pretreated withsulfuric acid to remove metals and followed with the exchange of Cl-forminto an OH-form. The effluent HA was analyzed with an ICP-MS method(Example 14). Results of the analysis were compared with initial HAsolution, the effluent through a cation-exchange resin (Rohm and HaasAmberjet 1500 (NH₄ ⁺)) and a 18 Megohm water, are shown in Table 6.

TABLE 6 Summary of ICP Analysis on Samples from Cation/Anion-ExchangeStudy Example 14 14 15 15 Sample 57% HA treated 57% HA purified 57% HAtreated 57% HA treated with a cation- with the anion- with a cation-with a cation- exchange resin Metal exchange resin exchange resinexchange resin (R&H 1500) and contents 18 Megohm of this invention (R&H1500 NH₄ ⁺)) (R&H 1500 NH₄ ⁺)) then with this (ppb) H₂O 57% HA only onlyfirst invention Al <0.20 40 15.45 9.71 Sb <0.10 <1.00 <1.00 As <0.10<1.00 <1.00 Ba <0.10 <1.00 <1.00 Be <0.10 <1.00 <1.00 Bi <0.10 <1.00<1.00 B 1.72 14.81 28.22 Cd <0.10 <1.00 <1.00 Ca <0.20 110 20.22 85.8761 <10 Cr <0.10 0.99 <1.00 <1.00 Co <0.10 <1.00 <1.00 Cu <0.10 99 19.1117.29 Ga <0.10 <1.00 <1.00 Ge <0.10 <1.00 <1.00 Au <0.10 <1.00 <1.00 Fe<0.20 140 <1.00 114.17 170 <10 Pb <0.10 84 <1.00 6.72 Li <0.20 <1.00<1.00 Mg <0.29 15 7.21 8.28 Mn <0.10 <0.5 <1.00 <1.00 Mo <0.10 <1.00<1.00 Ni <0.36 5.1 <1.00 5.19 Nb <0.10 <1.00 <1.00 K <0.64 330 139.8116.17 Si <0.50 <1.00 <1.00 Ag <0.10 <1.00 <1.00 Na 5.40 2500 1314.4243.12 <10 <10 Sr <0.50 <1.00 <1.00 Ta <0.10 <1.00 <1.00 Th <0.10 <1.00<1.00 Sn <0.10 0.78 <1.00 <1.00 Ti <0.10 10.99 9.18 V <0.20 <1.00 <1.00Zn 0.32 18 <1.00 32.59 Zr <0.10 <1.00 <1.00

Repeated addition of the complexation agent and the treatment of therestabilized HA with the anion-exchange resin can further improve thepurity of the HA free base. Combination of the treatment with NH4+-formof cation-exchange resin and the anion-exchange resin is illustratedalso in Table 6 (Example 15).

Example 16

In the manufacturing of HA free base, a coolant such as propylene glycolis used for the cooling of various process streams. Any leakage of thecoolant into the process-stream that contains a concentrated HA maycause undesired effects if metal impurities in the coolant are high. Acommercial propylene glycol was analyzed for several key metals and thecoolant was added with CDTA and treated with anion exchange resinfollowing the general procedure described in examples 14–15. Results ofmetal analysis on the initial coolant and samples purified by thepresent invention are illustrated in Table 7 below. As shown in thetable 7, the sample treated with a cation-exchange resin (PG(USP)/+exchange) still contain high levels of Cu, and Fe. Application of thisinvention lowers the metals to <10 and 140 ppb respectively for Cu andFe after three treatments (PG (USP)/+/−/−/− exchange).

TABLE 7 Metal Contents of a commercial Propylene Glycol and PurifiedSamples Sample PG(USP) as PG(USP)/ PG(USP)/ PG(USP)/ PG(USP)/ Metal(ppb) received + exchange +/− exchange +/−/− exchange +/−/−/− exchangeCa 8800  14 <10 <10  11 Cu 3600 1200 <10 <10 <10 Fe 22000  10000  780210 140 Ni  190  <30 <30 <30 <30 Na 4500  <10 <10 <10  36 Ti  <10  <10<10 <10 <10

Example 17

In another embodiment of the present invention is the addition ofstabilizer CDTA in the recycle stream to enhance the safety of themanufacturing process and the addition of the stabilizer in the productcontainer to improve the safety of product storage and shipping. Thestability of the hydroxylamine free base has been studied with anaccelerating rate calorimeter (ARC). For the ARC test, about 5 grams ofthe HA and HA with a stabilizer were introduced into a Pyrex glass testcell after the cell was washed with nitric acid and followed withrinsing by 18 Megohm deionized water. The test cell was sealed in air inthe ARC unit with a small headspace. The test cell was equipped with athermocouple contacting the bottom of the cell. A heater assembly wasprovided to head the surrounding of the sample cell and maintain thesurrounding of the cell at the same temperature of the sample cell.After a uniform temperature was established, a computer software programwas initiated for the heat-wait-search protocol. Once a self-heating ofthe sample was detected, the heating program was started to maintain anadiabatic condition for the test sample. Temperature and pressure datawere collected and the self-heat rate was determined.

A 57% HA free base having an iron metal content of 140 ppb wasinvestigated in an ARC unit. A similar HA solution was added with 100ppm of CDTA and investigated also with the same ARC unit. Results of theself-heat rate measurement were compared as shown in FIG. 5. This figureillustrates the enhanced stability of the concentrated HA by theaddition of a metal chelating agent, CDTA.

The addition of CDTA in the recycle stream will stabilize theconcentrated HA in the present process when it is produced. In addition,as the concentrated HA is produced, CDTA was added to enhance thestability of HA upon destined shipping and storage.

1. A process for manufacturing hydroxylamine comprising the following steps: (a) feeding an aqueous feed solution comprising hydroxylammonium ion and contaminants into a container containing at least one cation exchange resin to form a hydroxylammonium ion loaded resin; (b) feeding a water wash into said container to wash said loaded resin to remove said contaminants to form a purified hydroxylammonium ion loaded resin having said wash water associated therewith; (c) feeding an aqueous solution comprising hydroxylamine into said container to displace at least a portion of said associated wash water from said purified loaded resin to form a hydroxylamine solution associated resin; (d) feeding an aqueous desorbent into said container to desorb said hydroxylammonium ion from said associated resin to form an unbound hydroxylamine containing aqueous solution; (e) feeding a rinse water feed into said container to remove said unbound solution from said container to form a pure and concentrated aqueous solution comprising said hydroxylamine, and wherein step (c) is conducted prior to step (d).
 2. The process as defined in claim 1, wherein steps (a) through (e) are conducted at a column temperature ranging from 50 to 70° C.
 3. The process as defined in claim 1, wherein said aqueous feed solution comprises hydroxylamine sulfate.
 4. The process as defined in claim 1 wherein said resin is a sulfonated cation exchange resin.
 5. The process as defined in claim 1, wherein said desorbent is selected from the group consisting of a monovalent amine base; a monovalent hydroxide base and a mixture of the foregoing.
 6. The process as defined in claim 5, wherein the desorbent is aqueous ammonium hydroxide.
 7. The process as defined in claim 6, wherein said ammonium hydroxide has a ammonia concentration of at least 40% by weight.
 8. The protest as defined in claim 1, wherein said container is an ion exchange resin column and the superficial velocity of said aqueous desorbent and said rinse water added to said column in steps (d) and (a) is less than about 3.5 cm/min.
 9. The process as defined in claim 1, which further comprises treating said pure and concentrated solution with a metal chelating agent selected from the group consisting of cyclohexanediamminotetraacetic acid, thiourea, tetramethylthiuram disulfide, 4,8-dihydroxyquinoline-2-carboxylic acid, and a mixture of any of the foregoing agents.
 10. The, process as defined in claim 1, wherein during steps (a) through (e), further treating said container to cool said container with a coolant selected from a group consisting of propylene glycol, ethylene glycol, propylene glycol with water, ethylene glycol with water, and a mixture thereof to maintain the temperature of said container at 10 to 70° C.
 11. The process as defined in claim 10, wherein, prior to said cooling, said coolant is pretreated with a metal chelating agent selected from a group consisting of cyclohexanediamminotetraacetic acid, thiourea, 4,8-dihydroxyquinoline-2-carboxylic acid, and a mixture of any of the foregoing agents, to bind metal ion contaminants thereto and remove said contaminants from said coolant.
 12. The process as defined in claim 1 comprising the following steps: (a) feeding an aqueous feed solution including hydroxylammonium ion, hydrogen ions, ammonium ions, metal ions and anion con taminants into an ion exchange column containing at least one sulfonated cation exchange resin; (b) baiting the feeding of the aqueous feed solution and starting feeding of water into said ton exchange column when the ion exchange column effluent conductivity exceeds about 200 mMhos/cm to give a first ion exchange column water wash effluent stream; (c) halting the water feed of step (b) when the paid exchange column effluent conductivity reaches about 0 mMhos/cm; (d) feeding an aqueous hydroxylamine solution into said ion exchange column to displace essentially all of the water of said water feed from said column; (e) feeding an aqueous ammonium hydroxide solution having an ammonia concentration of at least 40% into the ion exchange column; (f) halting the feeding of said aqueous ammonium hydroxide solution into the ion exchange column when a sufficient amount has been added to completely desorb the hydroxylamine; (g) feeding a rinse water feed into the ion exchange column; (h) collecting a product stream emanating from the ion exchange column when the effluent density rises above about 1.085 g/cm³ wherein the product stream is an aqueous stream of essentially pure hydroxylamine and essentially no anion contaminants; (i) halting the collection of the product stream when effluent density decreases below about 1.075 g/cm³; and (j) halting the rinse water feed to the ion exchange column, and wherein step (d) is conducted prior to step (e).
 13. The process as defined in claim 12, therein steps (a) through (j) are conducted at a column temperature of 10 to 70° C.
 14. The process for manufacturing hydroxylamine of claim 12, wherein steps (a) through (j) are repeated in a cyclic process after the rinse water feed is halted in step (j).
 15. The process for manufacturing hydroxylamine of claim 14, wherein the process is accomplished in a continuous simulated moving bed ion exchange apparatus.
 16. The process as defined in claim 1 using a continuous simulated ion exchange apparatus including at least five ion exchange zones, each ion exchange column loaded with a sulfonated cation exchange resin, comprising the steps: (a) feeding an aqueous feed solution including a first concentration of hydroxylammonium ion, hydrogen ions, ammonium ions, metal ions and anion contaminants into a first ion exchange zone to give a sulfonated cation ion exchange resin loaded with hydroxylammonium ion; (b) feeding wish water into a second ion exchange zone including a sulfonated cation ion exchange resin loaded with hydroxylammonium ion to give a first ion exchange column water wash effluent stream; (c) feeding an aqueous hydroxylamine solution into a third ion exchange zone including a sulfonated cation ion exchange resin having water associated therewith to give a second ion exchange column effluent stream; (d) feeding an aqueous ammonium hydroxide solution having an ammonia concentration of at least 40% into a fourth ion exchange zone including a sulfonated cation ion exchange resin having hydroxylammonium ion associated therewith and collecting a third ion exchange zone effluent; (e) feeding a rinse water feed into a fifth ion; exchange zone to give a fourth ion exchange zone product stream; and (f) collecting the fourth ion exchange zone when the effluent density rises above about 1.085 g/cm³ and stopping the collecting of the fourth ion exchange ions effluent stream when the effluent density decreases to about 1.075 g/cm³ to give an aqueous hydroxylamine product having a second hydroxylamine concentration that is greater than the first hydroxylamine concentration in said aqueous feed solution and including essentially no anion contaminants.
 17. The process of claim 1, in which the hydroxylamine product is after-treated using an anion exchange resin. 